Multistage process for converting oxygenates to distillate hydrocarbons with interstage ethene recovery

ABSTRACT

An integrated process is provided for converting methanol, dimethylether or the like to heavy hydrocarbon products, especially distillate range hydrocarbons. In a first stage catalytic process oxygenate feedstock is converted to lower olefins. C 3   +  olefins are selectively sorbed in an interstage sorption fractionator and passed along with gasoline sorbent to a second stage oligomerization reactor. Distillate range hydrocarbons are useful as diesel fuel or the like.

FIELD OF THE INVENTION

This invention relates to a system for converting organic oxygenates, such as methanol or dimethyl ether (DME), to liquid hydrocarbons. In particular it provides a continuous process for producing distillate range fuel products by dehydrating and converting the oxygenate feedstock catalytically to produce an intermediate lower olefinic stream, and oligomerizing a fraction of the olefins to produce a major amount of distillate product for use as diesel fuel or the like.

BACKGROUND OF THE INVENTION

Recent developments in zeolite catalysts and hydrocarbon conversion processes have created interest in utilizing olefinic feedstocks, for producing C₅ ⁺ gasoline, diesel fuel, etc. In addition to the basic work derived from ZSM-5 type zeolite catalysts, a number of discoveries have contributed to the development of a new industrial process, known as Mobil Olefins to Gasoline/Distillate ("MOGD"). This process has significance as a safe, environmentally acceptable technique for utilizing feedstocks that contain lower olefins, especially C₂ -C₅ alkenes. This process may supplant conventional alkylation units. In U.S. Pat. Nos. 3,960,978 and 4,021,502, Plank, Rosinski and Givens disclose conversion of C₂ -C₅ olefins, alone or in admixture with paraffinic components, into higher hydrocarbons over crystalline zeolites having controlled acidity. Garwood et al have also contributed improved processing techniques to the MOGD system, as in U.S. Pat. Nos. 4,150,062, 4,211,640 and 4,227,992. The above-identified disclosures are incorporated herein by reference.

Conversion of lower olefins, especially propene and butenes, over HZSM-5 is effective at moderately elevated temperatures and pressures. The conversion products are sought as liquid fuels, especially the C₅ ⁺ aliphatic and aromatic hydrocarbons. Olefinic gasoline is produced in good yield by the MOGD process and may be recovered as a product or recycled to the reactor system for further conversion to distillate-range products. Operating details for typical MOGD units are disclosed in copending U.S. Pat. No. 4,456,779 (Owen et al) and U.S. Pat. No. 4,433,185 (Tabak), incorporated herein by reference.

In addition to their use as shape selective oligomerization catalysts, the medium pore ZSM-5 type catalysts are useful for converting methanol and other lower aliphatic alcohols and/or corresponding ethers to olefins. Particular interest has been directed to a catalytic process for converting low cost methanol to valuable hydrocarbons rich in ethene and C₃ ⁺ alkenes. Various processes are described in U.S. Pat. Nos. 3,894,107 (Butter et al), 3,928,483 (Chang et al), 4,025,571 (Lago), and in copending U.S. patent application Ser. No. 388,768, filed June 15, 1982 (Yurchak et al). Significance of the methanol-to-olefins ("MTO") type processes, especially for producing ethene, is discussed in Hydrocarbon Processing, November 1982, pp. 117-120. It is generally known that the MTO process can be optimized to produce a major fraction of C₂ -C₄ olefins.

SUMMARY OF THE INVENTION

It has been discovered that methanol, DME or the like may be converted to liquid fuels, particularly distillate, in a multi-stage continuous process, with a unique fractionation step between the major process conversion units.

In a preferred embodiment an integrated continuous process is provided for converting oxygenated organic feedstock to liquid hydrocarbons comprising methods and apparatus for contacting feedstock in a primary reactor stage with at least one catalyst comprising acidic zeolite at elevated temperature and moderate pressure to dehydrate and convert at least a portion of the feedstock to hydrocarbons containing a major fraction of C₂ -C₄ olefins and a minor fraction containing C₅ ⁺ heavy hydrocarbon. After cooling and separating primary stage effluent to provide an aqueous liquid stream, a heavy hydrocarbon liquid stream and a light hydrocarbon vapor stream rich in C₂ -C₄ olefins, the light hydrocarbon stream is fractionated by selectively sorbing C₃ ⁺ hydrocarbons in a gasoline sorbent stream to recover an ethene-rich vapor stream and a liquid stream rich in C₃ ⁺ sorbate. In a secondary reactor stage the sorbate-rich stream is contacted with oligomerization catalyst comprising medium-pore shape selective acidic zeolite at substantially increased pressure and elevated temperature to convert olefins to a heavier liquid hydrocarbon effluent stream comprising olefinic gasoline and distillate range liquids. The liquid hydrocarbon effluent stream from the secondary stage is fractionated to obtain a distillate stream, gasoline stream and lighter hydrocarbon stream, at least a portion of the gasoline stream is recycled as sorbent. The oligomerization catalyst preferably comprises ZSM-5 type zeolite.

Other objects and features of the invention will be seen in the following description and drawings.

THE DRAWINGS

FIG. 1 is a process flow sheet showing the major unit operations and process streams;

FIG. 2 is a schematic representation of a preferred multi-reactor system and fractionation system for dehydration and conversion of methanol and/or DME to lower olefins;

FIG. 3 is a preferred sorption fractionation system; and

FIG. 4 is a typical olefin conversion reactor system for distillate mode operation.

DESCRIPTION OF PREFERRED EMBODIMENTS

Catalyst versatility permits the same zeolite to be used in both the primary stage (MTO) and secondary oligomerization stage (MOGD). While it is within the inventive concept to employ substantially different catalysts in these stages, it is advantageous to employ a standard ZSM-5 having a silica alumina molar ratio of 70:1.

The oligomerization catalysts preferred for use herein include the medium pore shape selective crystalline aluminosilicate zeolites having a silica to alumina ratio of at least 12, a constraint index of about 1 to 12 and acid cracking activity of about 160-200. Representative of the ZSM-5 type zeolites are ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38. ZSM-5 is disclosed in U.S. Pat. No. 3,702,886 and U.S. Pat. No. Re. 29,948. Other suitable zeolites are disclosed in U.S. Pat. Nos. 3,709,979; 3,832,449, 4,076,979, 3,832,449, 4,076,842, 4,016,245 and 4,046,839. The disclosures of these patents are incorporated herein by reference. A suitable catalyst for each fixed bed operation consisting essentially of HZSM-5 zeolite with 35 wt.% alumina binder in the form of cyclindrical extrudates of about 1-5 mm diameter. Other catalysts which may be employed for converting methanol/DME to lower olefins include the borosilicate, ferrosilicate, "silicalite" and/or synthetic mordenite materials.

In this description, metric units and parts by weight are employed unless otherwise stated. While various reactor configurations may be used, including fluidized bed catalytic reactors, moving bed and fixed bed reactors, the invention is described for use in a plurality of fixed bed reactors operated under differing process conditions depending upon relative position in the system.

Referring to FIG. 1, oxygenate feedstock (methanol or DME, for instance) is fed to the primary stage where it is converted to a lower olefin and heavier hydrocarbon plus water by dehydration of the oxygenated feedstock. Byproduct water is recovered by simple phase separation from the cooled effluent. Liquid hydrocarbons (HC) comprising C₅ ⁺ gasoline range materials may be recovered or, optionally combined with the lower olefinic vapor for further conversion in the secondary stage. The olefinic vapor phase effluent from the primary stage is then contacted with a gasoline sorbent stream in the sorption fractionation section to provide a C₃ ⁺ liquid stream, which is passed to the oligomerization reaction section. The combined olefinic stream (containing recycled olefinic gasoline rich in C₃ -C₄ olefin sorbate) is reacted at high pressure and elevated temperature over the oligomerization catalyst. Secondary stage effluent is then separated into light gases, C₅ ⁺ gasoline for recycle and distillate range hydrocarbons. The distillate stream contains a major fraction of high boiling aliphatics and a minor amount of aromatics. The distillate product may be further stabilized by hydrotreating (HDT) in a relatively mild process to saturate the olefinic compounds and convert aromatics to corresponding naphthenes without substantial cracking or dealkylation to yield a distillate fuel product. Ethylene (ethene, C₂ H₂) is recovered as a valuable chemical feedstock from the process.

In the process for catalytic conversion of olefins to heavier hydrocarbons by catalytic oligomerization using an acid crystalline zeolite, such as ZSM-5 type catalyst, process conditions can be varied to favor the formation of either gasoline or distillate range products. At moderate temperature and relatively high pressure, the conversion conditions favor distillate range product having a normal boiling point of at least 165° C. (330° F.). Lower olefinic feedstocks containing C₂ -C₆ alkenes may be converted selectively; however, the distillate mode conditions do not convert a major fraction of ethylene. While propene, butene-1 and others may be converted to the extent of 50 to 95% in the distillate mode, only about 10 to 50% of the ethylene component will be consumed. Accordingly, the ethene is advantageously recovered prior to the oligomerization stage.

Primary stage conversion operation in FIG. 2 depicts a fixed bed multi-reactor system for converting methanol (CH₃ OH) and/or DME. A typical crude methanol feedstock may contain 4 to 17% water, with minor amounts of carbon oxides, methane, DME, etc.

In the primary stage, ethene production may be optimized by employing fixed bed primary stage conditions in the temperature range of about 260° C. to 425° C., a pressure range of about 170 to 800 kPa and weight hourly space velocity range of about 0.5 to 1.0 based on ZSM-5 equivalent catalyst and methanol equivalent in the primary stage feedstock. Typically about 25 to 90% of MeOH/DME feedstock is converted per reactor pass and water diluent is cofed with methanol and/or dimethyl ether in a molar ratio of about 0.1:1 to 5:1. Under these conditions, the primary stage hydrocarbon effluent usually contains about 25 to 40 wt.% ethene, about 10 to 50 wt.% propene, about 2 to 30 wt.% butene, less than 10 wt.% C₁ to C₄ paraffins, and about 5 to 20 wt.% aromatics, including about 1 to 5 wt.% durene.

In the preferred embodiment of FIG. 2, the feedstock is methanol (MeOH), which may be partially dehydrated in a separate process step over gamma-alumina catalyst to yield dimethyl ether (DME) and water. A preliminary dewatering step can be used to provide a feedstock consisting essentially of MeOH and/or DME; however, the presence of water in the MTO reactor may be beneficial. The feedstock is fed continuously under low pressure through line 1 and heat exchanger 2, where it is raised to process temperature, and introduced to the first stage MTO reactor system. The initial dehydration reactor 10 is followed by a series of fixed bed catalytic reactors 11, 12, 13, 14 and 15 containing zeolite conversion catalyst. Inter-reactor coolers 10A, 11A, 12A, 13A, 14A and effluent cooler 15A control temperature of the process stream. Interstage separation means is provided for recovering water and hydrocarbons from the primary stage effluent stream. The first stage effluent is cooled to condense water and a major amount of C₅ ⁺ liquid hydrocarbons. These liquids are separated from the hydrocarbon vapor in phase separator means 16. Byproduct water may be recovered from unreacted feedstock and discarded, or optionally a portion may be recycled, as indicated by the dashed lines. The liquid hydrocarbon phase and the ethene-rich light hydrocarbon vapor streams are recovered from separator 16. Optionally unconverted DME may be removed by absorber 20. This optional section includes a compressor 19 and stripper 21. Other suitable fixed bed catalytic processes for conversion of methanol/DME to lower olefins are described in U.S. Pat. Nos. 4,387,263, 4,393,265, 4,361,715 and South African Patent Application Ref. No. V01750 (Clover et al) filed Mar. 30, 1983, the entire disclosures of which are incorporated herein by reference. While the primary stage dehydration reactor has been exemplified herein by a fixed bed unit, a suitable fluid catalyst apparatus is disclosed in U.S. Pat. No. 4,379,123 (Daviduk and Haddad).

A typical MTO operation is conducted over a fixed bed of small crystal (0.02-0.05μ) HZSM-5/alumina extrudate catalyst at about 170 kPa (25 psia), with a 1:1 H₂ O:CH₃ OH equivalent ratio at 315° C. (600° F.) at a space velocity (WHSV=0.5-1) to convert about 50% of the oxygenated organic feedstock components to hydrocarbons. Table I lists the organic hydrocarbon product distribution from a typical MTO process.

                  TABLE I                                                          ______________________________________                                         MTO Product Distribution                                                       Component        wt. %                                                         ______________________________________                                         Methane, wt. %   0.6                                                           Ethylene, wt. %  26.2                                                          Ethane, wt. %    0.1                                                           Propylene, wt. % 22.8                                                          Propane          3.9                                                           Butenes          7.9                                                           Isobutane        3.9                                                           n-Butane         2.6                                                           Pentenes         2.4                                                           C.sub.5 P + N    7.1                                                           C.sub.6 P + N    5.1                                                           C.sub.7 O        0.6                                                           C.sub.7 P + N    3.2                                                           C.sub.7 O        0.7                                                           C.sub.8 P + O + N                                                                               2.1                                                           C.sub.9 P + O + N                                                                               1.3                                                           C.sub.10 P + O + N                                                                              1.1                                                           Benzene          0.1                                                           Toluene          0.5                                                           C.sub.8 Aromatics                                                                               3.5                                                           C.sub.9 Aromatics                                                                               2.1                                                           C.sub.10 Aromatics                                                                              2.2                                                           (Durene)         (1.7)                                                         ______________________________________                                    

SORPTION FRACTIONATION

The light hydrocarbon stream recovered from the primary conversion stage preferably contains a major amount of C₂ -C₄ olefins. The novel system includes a first fractionation means for recovering ethene from the primary stage olefinic vapor including a sorption tower operatively connected to selectively sorb C₃ ⁺ hydrocarbons from the olefinic vapor in a liquid sorption stream. Since the interstage fractionation unit is usually operated at a pressure higher than the primary stage and lower than the secondary conversion stage, vapor compression means for the primary stage light hydrocarbon stream and means for pressurizing and heating the liquid sorption stream containing C₃ ⁺ sorbate are provided.

A suitable sorption fractionation system is described in copending U.S. patent application Ser. No. 508,779 filed June 29, 1983 (Hsia et al), the disclosure of which is incorporated herein by reference. The C₂ ⁻ and C₃ ⁺ separation is accomplished by a single absorber-stripper using gasoline recycle as absorbent and pumparounds for removing absorption heat. The amount of absorbent is set by the amount of recycle gasoline required in the C₃ ⁺ olefins conversion reaction thereby allowing the tower bottom stream to be pumped directly to the reactor pressure. Without using refrigeration, this tower efficiently and effectively separates the ethylene and light gases (H₂, CO, CO₂ and CH₄) from the C₃ ⁺ hydrocarbon.

The gasoline sorbent is an aliphatic hydrocarbon mixture boiling in the normal gasoline range of about 50° to 165° C. (125° to 330° F.), with minor amounts of C₄ -C₅ alkanes and alkenes. Preferably, the total gasoline sorbent stream to feedstock weight ratio is greater than about 3:1; however, the content of C₃ ⁺ olefinic components in the feedstock is a more preferred measure of sorbate to sorbent ratio. Accordingly, the process may be operated with a mole ratio of about 0.2 moles to about 10 moles of gasoline per mole of C₃ ⁺ hydrocarbons in the feedstock, with optimum operation utilizing a sorbent:sorbate molar ratio about 1:1 to 1.5:1.

It is understood that the various process conditions are given for a continuous system operating at steady state, and that substantial variations in the process are possible within the inventive concept. In the detailed examples, metric units and parts by weight are employed unless otherwise specified.

Referring to FIG. 3 of the drawing, olefinic feedstock is introduced to the system through a feedstock inlet 101, passing through knockout pot 102 to remove any condensate, such as water. At this point the vapor is at a temperature from ambient up to about 38° C. (100° F.) and a pressure of about 60 kPa (75 psig), and is pressurized by compressor means 103 to about 2230 kPa (310 psig). This pressurized stream is then mixed with a liquid downcomer stream 104 from fractionating tower 120, cooled in exchanger 105 to about 35°-40° C. and passed to phase separator 106 where condensed water is removed. The hydrocarbon stream is then fed via conduit 107 connected between stages of fractionating sorption tower 120 wherein gaseous olefinic feedstock is contacted with liquid sorbent in a vertical fractionation column operating at least in the upper portion thereof in countercurrent flow. Effectively this unit is a C₂ /C₃ ⁺ splitter. The sorption tower employs a plate column; however, the fractionation equipment may employ vapor-liquid contact means of various designs in each stage including packed beds of Raschig rings, saddles or other porous solids or low pressure drop contact devices.

Sorption tower 120, as depicted, has multiple contact zones, with the heat of absorption being removed via interstage pump around cooling means 121A, B, and C. The liquid gasoline sorbent is introduced to the sorption tower through an upper inlet means 125 above the top contact section 120T. It is preferred to mix incoming liquid sorbent with outgoing splitter overhead ethylene-rich gas. High purity ethylene is recovered from the system through gas outlet 131 and sent to storage, further processing or conversion to other products. Liquid sorbent from separator 130 is then pumped to the upper liquid inlet 124 for countercurrent contact in a plate column or the like with upwardly flowing ethylene rich vapors. Liquid from the bottom of upper contact zone is pumped to a heat exchanger in loop 121A, cooled and returned to the tower in loop 121B adjacent an intermediate contact zone, again cooled in loop 121C, and returned to the tower above contact zone 126, which is located at the feedstock inlet 1. Under tower design conditions of about 2060 kPa (300 psia), it is preferred to maintain liquid temperature of streams entering the tower from 121, 122 and 124 at about 40° C. (100° F.). The lower contact zone provides further fractionation of the olefin-rich liquid. Heat is supplied to the sorption tower by removing liquid from the bottom via reboiler loop 128, heating this stream in a heat exchanger, and returning the reboiled bottom stream to the tower below contact zone 126. The liquid sorbate-sorbent mixture is withdrawn through bottom outlet 129 and pumped to storage or directly to the secondary stage for further reaction. The fractionator bottoms stream 129 is recovered at about 120° C. (250° F.), then pumped to the higher reactor pressure (e.g., about 4670 kPa, 665 psig) and passed to the secondary conversion stage. A typical sorption fractionation material balance is given below, for steady state operation using olefinic feed gas from the MTO primary stage. The units are expressed in moles per hour.

                  TABLE II                                                         ______________________________________                                         SORPTION FRACTIONATION MATERIAL BALANCE                                                              GAS         LIQUID                                       FEED      RECYCLE     PRODUCT     SORBATE                                      GAS       GASOLINE    STREAM      STREAM                                       ______________________________________                                         H.sub.2                                                                             1.17     --          1.17      --                                         CO   0.22     --          0.22      --                                         CO.sub.2                                                                            0.38     --          0.38      --                                         C.sub.1                                                                             0.63     --          0.63      --                                         C.sub.2                                                                             14.83    --          14.78     0.05                                       C.sub.2                                                                             0.11     --          0.10      0.01                                       C.sub.3                                                                             8.09     --          0.36      7.73                                       C.sub.3                                                                             1.62     --          0.02      1.60                                       iC.sub.4                                                                            1.00     0.09        0.03      1.06                                       C.sub.4                                                                             2.40     0.03        0.01      2.42                                       nC.sub.4                                                                            0.32     0.17        0.04      0.45                                       iC.sub.5                                                                            0.67     2.64        0.25      3.06                                       C.sub.5                                                                             0.24     1.23        0.11      1.36                                       nC.sub.5                                                                            0.06     0.22        0.02      0.26                                       C.sub.6.sup.+                                                                       0.81     11.19       0.13      11.87                                      H.sub.2 O                                                                           0.37     --          0.03      --                                              32.92    15.58       18.27     29.88                                      ______________________________________                                    

SECONDARY STAGE Oligomerization Reactor Operation

The secondary stage provides catalytic oligomerization reactor means containing medium pore shape selective zeolite oligomerization catalyst for converting olefinic hydrocarbons in the sorption stream to liquid hydrocarbons comprising a major amount of distillate. This process stream is passed to second fractionation means for separating secondary stage effluent into a light hydrocarbon stream rich in LPG (C₃ -C₄) aliphatic hydrocarbons, a C₅ ⁺ gasoline stream and distillate range stream. Means is provided for recycling at least a portion of the C₅ ⁺ gasoline stream to the first sorption fractionation means as a lean sorbent stream.

A typical distillate mode secondary stage reactor system 220 is shown in FIG. 4. A plural reactor system may be employed with inter-reactor cooling, whereby the reaction exotherm can be carefully controlled to prevent excessive temperature above the normal moderate range of about 190° to 315° (375°-600° F.). The olefinic feedstream comprising the C₃ ⁺ light hydrocarbons and sorbent gasoline is introduced through conduit 229 and carried by a series of conduits through heat exchangers 212A, B, C and furnace 214 where the feedstream is heated to reaction temperature. The olefinic feedstream is then carried sequentially through a series of zeolite beds 220A, B, C wherein at least a portion of the olefin content is converted to heavier distillate constituents. Advantageously, the maximum temperature differential across only one reactor is about 30° C. (ΔT˜50° F.) and the space velocity (LHSV based on olefin feed) is about 0.5 to 1.5. The heat exchangers 212A and 212B provide inter-reactor cooling and 212C reduces the effluent to separation temperature. An optional heat exchanger 212D may further recover heat from the effluent stream 121 prior to separation.

Preferably, the secondary stage process conditions are optimized to produce heavy liquid hydrocarbons having a normal boiling above 165° C. (330° F.) are fed as a continuous stream to a final fractionator unit (such as a distillation system). Gasoline, rich in C₅ ⁺ olefins and lighter hydrocarbons are fractionated in a tower to provide an olefinic gasoline stream for recycle to the MOGD reactor system or recovered as product. The lighter hydrocarbons, rich in C₃ -C₄ alkenes may be condensed and recovered as LPG product or optionally recycled to the MOGD reactor system. The secondary stage typical HZSM-5 fixed bed reactor system depicted in FIG. 4, operates at 0.6 liquid hourly space velocity (based on olefins fed to reactors), 1:1 gasoline:olefin recycle ratio, temperature of 230° C. (450° F.) (SOC) to 315° C. (600° F.) (EOC) and a total pressure of 4225 kPa (600 psig) at minimum olefin partial pressure at the inlet of 1100 kPa (160 psig). The secondary stage effluent from such a typical system is shown below.

                  TABLE III                                                        ______________________________________                                         Component         Wt. %                                                        ______________________________________                                         CH.sub.4          0.07                                                         C.sub.2 H.sub.6   0.13                                                         C.sub.3 H.sub.8   2.80                                                         IC.sub.4 H.sub.10 2.00                                                         NC.sub.4 H.sub.10 2.00                                                         i-C.sub.5 H.sub.12                                                                               0.42                                                         n-C.sub.5 H.sub.12                                                                               0.03                                                         C.sub.5 H.sub.10  0.95                                                         C.sub.6 -330° (Gasoline)                                                                  12.60                                                        330° + (Distillate)                                                                       79.00                                                        ______________________________________                                    

A typical MOGD secondary stage fractionation system is described in copending U.S. patent application Ser. No. 488,834 filed Apr. 26, 1983 (Owen et al), incorporated herein by reference.

It is within the inventive concept to cascade substantially all C₃ ⁺ vapor and liquid hydrocarbon first stage product into an MOGD reactor followed by hydrotreating of the distillate product as depicted in the dashed lines of FIG. 1. This will minimize the number of process steps and will maximize distillate production by polymerizing gasoline range olefins, and by alkylating gasoline range aromatics. Durene can be reduced via saturation to its corresponding naphthene is a subsequent mild hydrotreating step. Substantially all of the polymethylbenzenes or other aromatics formed in the dehydration reactor stage can be accumulated in the distillate fraction according to the present invention, and hydrotreated distillate durene content is decreased substantially below 2 wt.%, preferably below 1%.

The present process is particularly useful in producing a major product stream wherein the 165 C⁺ fraction consists mainly of C₁₀ to C₂₀ aliphatic hydrocarbons containing a minor amount of cyclic components. The low temperature, high pressure distillate mode secondary stage operation favors the formation of linear oligomers. The preferred distillate mode operation can provide a larger part of the total (non-ethene) fuel products as heavy hydrocarbons, while retaining a high yield of valuable ethene product. 

I claim:
 1. An integrated continuous process for converting oxygenated organic feedstock to liquid hydrocarbons comprising the steps of contacting feedstock in a primary reactor stage with at least one catalyst comprising acidic zeolite catalyst at elevated temperature and moderate pressure to dehydrate and convert at least a portion of the feedstock to hydrocarbons containing a major fraction of C₂ -C₄ olefins and a minor fraction containing C₅ ⁺ heavy hydrocarbons;cooling and separating primary stage effluent to provide an aqueous liquid stream, a heavy hydrocarbon liquid stream and a light hydrocarbon vapor stream rich in C₂ -C₄ olefins; fractionating the light hydrocarbon stream by selectively sorbing C₃ ⁺ hydrocarbons in a gasoline sorbent stream to recover an ethene-rich vapor stream and a liquid stream rich in C₃ ⁺ sorbate; contacting the sorbate-rich stream in a secondary reactor stage with oligomerization catalyst comprising medium-pore shape selective acidic zeolite at substantially increased pressure and elevated temperature to convert olefins to a heavier liquid hydrocarbon effluent stream comprising olefinic gasoline and distillate range liquids; fractionating the liquid hydrocarbon effluent stream from the secondary stage to obtain a distillate stream, gasoline stream and lighter hydrocarbon stream; and recycling at least a portion of the gasoline stream as sorbent.
 2. The process of claim 1 wherein recycled gasoline comprising C₅ ⁺ olefinic components is further reacted in the secondary stage to produce heavier hydrocarbon.
 3. The process of claim 1 wherein the oligomerization catalyst comprises HZSM-5 type zeolite.
 4. THe process of claim 1 wherein primary stage feedstock comprising methanol and/or dimethyl ether is converted over HZSM-5 type catalyst to hydrocarbons comprising a major amount of C₂ -C₄ olefins and a minor amount of normally liquid hydrocarbons.
 5. The process of claim 4 wherein ethene production is optimized by employing fixed bed primary stage conditions in the temperature range of about 260° C. to 425° C., a pressure range of about 170 to 800 kPa and weight hourly space velocity range of about 0.5 to 1.0 based on ZSM-5 equivalent catalyst and methanol equivalent in the primary stage feedstock.
 6. The process of claim 5 wherein about 25 to 90% of methanol feedstock is converted per reactor pass and wherein water diluent is cofed with methanol and/or dimethyl ether in a molar ratio of about 0.1:1 to 5:1,based on methanol equivalents.
 7. The process of claim 6 wherein primary stage hydrocarbon effluent contains about 5 to 40 wt.% ethene, about 10 to 60% wt.% propene, about 2 to 30 wt.% butene, less than 10 wt.% C₁ to C₄ paraffins, about 5 to 20 wt.% aromatics, including about 1 to 5 wt.% durene.
 8. The process of claim 1 wherein the minor fraction of C₅ ⁺ heavy hydrocarbons is combined with the sorbate-rich stream in the secondary reactor stage.
 9. The process of claim 1 wherein the secondary stage process conditions are optimized to produce a major amount of heavy liquid hydrocarbons having a normal boiling point greater than about 165° C.
 10. The process of claim 9 wherein the secondary stage employs a fixed bed of shape selective medium pore zeolite catalyst to oligomerize olefins at a temperature of about 190° C. to 315° C. and total pressure of about 4200 to 7000 kPa, said zeolite having a silica:alumina molar ratio greater than about 12:1 and a constraint index of about 1 to
 12. 11. An integrated continuous process for converting feedstock comprising methanol, dimethylether or mixtures thereof to liquid hydrocarbons comprising the steps of contacting feedstock in a primary reactor stage with zeolite catalyst at elevated temperature and moderate pressure to dehydrate and convert at least a portion of the feedstock to light hydrocarbons comprising C₂ -C₄ olefins and by-product water;cooling and separating primary stage effluent to provide an aqueous liquid stream and a light hydrocarbon vapor stream rich in C₂ -C₄ olefins; fractionating the light hydrocarbon stream by selectively sorbing C₃ ⁺ hydrocarbons in a liquid hydrocarbon sorbent stream to recover an ethene-rich vapor stream and a liquid stream rich in C₃ ⁺ sorbate; contacting the sorbate-rich stream in a secondary reactor stage with oligomerization catalyst comprising medium-pore shape selective acidic zeolite at substantially increased pressure and elevated temperature to convert olefins to a heavier liquid hydrocarbon effluent stream consisting essentially of olefinic liquids; and recycling a fraction of the olefinic liquid stream as sorbent.
 12. The process of claim 11 wherein the oligomerization catalyst comprises HZSM-5 type zeolite. 